Linked perfusion to continuous-flow stirred-tank reactor cell culture system

ABSTRACT

Methods of protein production in a linked culture and production bioreactor system are provided. Such methods include a culture bioreactor (N-1 bioreactor) linked to production bioreactor (N bioreactor). More specifically, the methods include (a) culturing cells with a gene that encodes the protein of interest in a continuous perfusion culture bioreactor (N-1 bioreactor); inoculating a continuously stirred tank reactor (CSTR) production bioreactor (N bioreactor) with cells obtained from step (a); and culturing the cells in the CSTR production bioreactor under conditions that allow production of the protein of interest.

FIELD OF THE SUBJECT TECHNOLOGY

The subject technology relates to methods of protein production incultured animal cells, preferably mammalian cells, using a linkedbioreactor system including a culture bioreactor (N-1 bioreactor) linkedto production bioreactor (N bioreactor). More specifically, the subjecttechnology relates to a continuous perfusion culture bioreactor linkedto a chemostat or continuous-flow stirred-tank reactor (CSTR) productionbioreactor.

BACKGROUND OF THE SUBJECT TECHNOLOGY

Conventional perfusion cell culture systems suffer from disadvantagessuch as large-volume media consumption, long times to reach peak celldensities and complications with cell retention devices when used at thelargest scale.

For example, in the hybrid cell culture systems disclosed in WO2008091113, multiple CSTR bioreactors are linked in series, each usingcell retention devices (e.g., packed bed). The reference suggests thatin an ideal system, the last bioreactor in the series would use completecell retention, but that the earlier bioreactors would allow some cellsto pass out of the bioreactor. In the hybrid cell culture systemdisclosed in WO 2015003012, a single ‘nurse’ or N-1 bioreactorperiodically inoculates multiple N-stage production bioreactors whichhave cell retention devices. In the hybrid cell culture system disclosedin WO 2015095809, an N-1 perfusion bioreactor is used to prepare aninoculum for a production bioreactor operating in a fed-batch orperfusion mode. However, all the data and examples in this referencesuggest that the N-1 is a perfusion bioreactor operating for a shortperiod of time, producing a single inoculum used to inoculate aproduction bioreactor which operates as a fed-batch. Additionally,nowhere do the authors of this application suggest that the productionbioreactor might operate as a chemostat (or CSTR) with no cell retentionthat produces a continuous harvest to the downstream purificationoperation.

Cell retention systems are difficult to operate and design for use atthe large scale (>1,000 L). For example, nearly any cell retentiondevice that uses a membrane (as do many that are currently in use atlarge scale) will eventually plug with cell debris. This plugging ismore likely to occur with low cell viability cultures as the cells aremore fragile, and the particulates that are formed are often similar insize to the pores of the membrane (0.2-5 micron). Additionally, asmembranes plug with cell debris they also begin to effectively functionas ultrafiltration devices, retaining the high molecular weight productproteins within the bioreactor in a not easily predictable orreproducible manner. This is a disadvantage because in a continuousperfusion bioreactor system it is advantageous to have the product ofinterest be continuously removed from the bioreactor in the cell-freeharvest and delivered to the downstream operation in a consistentmanner.

As a result, conventional continuous perfusion cell culture systemsusually have working volumes below 2,000 L, and if operated underconditions where productivity is highest (e.g. high viable celldensities and high perfusion rates), require frequent change out of themembrane based cell retention device due to plugging and productretention (ultrafiltration). Additionally, cell retention devicescapable of handling very large volumes of cell-free culture harvest canbe quite complex (e.g. many moving parts) and expensive, damaging tocells through excessive shear forces, and prone to failure. As the cellretention devices are typically external to the bioreactor, effectivecleaning and sterilization can also be challenging at the large scale.These factors can result in high cost of operation, significant loss ofproductivity and inefficiencies in production of a protein of interest.Therefore, there still remains a need for an alternative cell culturemethod or system that overcomes the limitations associated with thecurrent conventional perfusion culture systems.

SUMMARY OF THE SUBJECT TECHNOLOGY

The subject technology is illustrated, for example, according to variousaspects and embodiments listed below.

In one aspect, the subject technology relates to a method of producing aprotein of interest, including: (a) culturing cells comprising a genethat encodes the protein of interest in a culture bioreactor (N-1bioreactor); (b) inoculating a production bioreactor (N bioreactor) withcells obtained from step (a); and (c) culturing the cells in theproduction bioreactor under conditions that allow production of theprotein of interest. In one or more embodiments relating directly orindirectly to this aspect, the method further comprises step (d)harvesting the protein of interest from the production bioreactor; theculture bioreactor is a continuous perfusion culture bioreactor and theproduction bioreactor is a continuously stirred tank reactor (CSTR)production bioreactor; the production bioreactor has no cell retentiondevice; volume ratio of the culture bioreactor to the productionbioreactor is about 1:1 to about 1:20; volume ratio of the culturebioreactor to the production bioreactor is about 1:1 to about 1:5;volume ratio of the culture bioreactor to the production bioreactor isabout 1:5; the inoculation in step (b) is by transferring cells from theculture bioreactor to the production bioreactor; the cell transfer is bycell bleed in continuous or semi-continuous modes; the cell transfer isin semi-continuous mode including the cell transfer once at every periodof time from 2 minutes to 24 hours or at any interval therebetween; step(a) optionally alternates between a first and second culture bioreactorsto allow for renewal and continuous production of culture cells for usein step (b); the second culture bioreactor is a continuous perfusionculture bioreactor; the production bioreactor operates continuously fora period of greater than 3 weeks; the production bioreactor operatescontinuously for a period of greater than 4 weeks; the productionbioreactor operates continuously for a period of greater than 5 weeks;the production bioreactor operates continuously for a period of greaterthan 6 weeks; harvesting step (d) is continuous; the cells are CHOcells, HEK-293 cells, VERO cells, NSO cells, PER.C6 cells, Sp2/0 cells,BHK cells, MDCK cells, MDBK cells or COS cells; the productionbioreactor has a volumetric productivity of at least 0.6 grams per literper day for a period of at least 14 days; the production bioreactor hasa volumetric productivity of at least 0.6 grams per liter per day for aperiod of at least 20 days; the production bioreactor has a volumetricproductivity of at least 0.6 grams per liter per day for a period of atleast 30 days; the production bioreactor has a product residence time ofabout 1 to about 10 days; the production bioreactor has a dilution rateof about 1 to about 0.1 volume per day; the production bioreactor is fedwith a diluent solution; the diluent solution is water or saline.

In another aspect, the subject technology relates to a linked cultureprocess, including: (a) culturing cells comprising a gene that encodesthe protein of interest in a culture bioreactor (N-1 bioreactor); (b)inoculating a production bioreactor (N bioreactor) with cells obtainedfrom step (a); and (c) culturing the cells in the production bioreactorunder conditions that allow production of the protein of interest. Inone or more embodiments relating directly or indirectly to this aspect,the linked culture process further includes step (d) harvesting theprotein of interest from the production bioreactor; the culturebioreactor is a continuous perfusion culture bioreactor and thecontinuous production bioreactor is a continuously stirred tank reactor(CSTR) production bioreactor; the production bioreactor has no cellretention device; volume ratio of the culture bioreactor to theproduction bioreactor is about 1:1 to about 1:20; volume ratio of theculture bioreactor to the production bioreactor is about 1:1 to about1:5; volume ratio of the culture bioreactor to the production bioreactoris about 1:5; the inoculation in step (b) is by transferring cells fromthe culture bioreactor to the production bioreactor; the cell transferis by cell bleed in continuous or semi-continuous modes; the celltransfer is in semi-continuous mode comprising the cell transfer once atevery period of time from 2 minutes to 24 hours or at any intervaltherebetween; step (a) optionally alternates between a first and secondculture bioreactors to allow for renewal and continuous production ofculture cells for use in step (b); the second culture bioreactor is acontinuous perfusion culture bioreactor; the production bioreactoroperates continuously for a period of greater than 3 weeks; theproduction bioreactor operates continuously for a period of greater than4 weeks; the production bioreactor operates continuously for a period ofgreater than 5 weeks; the production bioreactor operates continuouslyfor a period of greater than 6 weeks; harvesting step (d) is continuous;the cells are CHO cells, HEK-293 cells, VERO cells, NSO cells, PER.C6cells, Sp2/0 cells, BHK cells, MDCK cells, MDBK cells or COS cells; theproduction bioreactor has a volumetric productivity of at least 0.6grams per liter per day for a period of at least 14 days; the productionbioreactor has a volumetric productivity of at least 0.6 grams per literper day for a period of at least 20 days; the production bioreactor hasa volumetric productivity of at least 0.6 grams per liter per day for aperiod of at least 30 days; the production bioreactor has a productresidence time of about 1 to about 10 days; the production bioreactorhas a dilution rate of about 1 to about 0.1 volume per day; theproduction bioreactor is fed with a diluent solution; the diluentsolution is water or saline.

In another aspect, the subject technology relates to a linked culturesystem, including: a culture bioreactor (N-1 bioreactor) for culturingcells comprising a gene that encodes the protein of interest; and aproduction bioreactor (N bioreactor) which is linked to the culturebioreactor and receives cells as inoculum from the culture bioreactor,wherein the production bioreactor operates under conditions that allowproduction of the protein of interest. In one or more embodimentsrelating directly or indirectly to this aspect, the system is linked toa downstream purification system for harvesting the protein of interestfrom the production bioreactor; the culture bioreactor is a continuousperfusion culture bioreactor and the production bioreactor is acontinuously stirred tank reactor (CSTR) production bioreactor; theproduction bioreactor has no cell retention device; volume ratio of theculture bioreactor to the production bioreactor is about 1:1 to about1:20; volume ratio of the culture bioreactor to the productionbioreactor is about 1:1 to about 1:5; volume ratio of the culturebioreactor to the production bioreactor is about 1:5; the inoculation isby transferring cells from the culture bioreactor to the productionbioreactor; the cell transfer is by cell bleed in continuous orsemi-continuous modes; the cell transfer is in semi-continuous modecomprising the cell transfer once at every period of time from 2 minutesto 24 hours or at any interval therebetween; the system includes asecond culture bioreactor that operates in parallel with a first culturebioreactor and alternates in producing inoculum for transfer to theproduction bioreactor; the second culture bioreactor is a continuousperfusion culture bioreactor; the production bioreactor operatescontinuously for a period of greater than 3 weeks; the productionbioreactor operates continuously for a period of greater than 4 weeks;the production bioreactor operates continuously for a period of greaterthan 5 weeks; the production bioreactor operates continuously for aperiod of greater than 6 weeks; the harvesting is continuous; the cellsare CHO cells, HEK-293 cells, VERO cells, NSO cells, PER.C6 cells, Sp2/0cells, BHK cells, MDCK cells, MDBK cells or COS cells; the productionbioreactor has a volumetric productivity of at least 0.6 grams per literper day for a period of at least 14 days; the production bioreactor hasa volumetric productivity of at least 0.6 grams per liter per day for aperiod of at least 20 days; the production bioreactor has a volumetricproductivity of at least 0.6 grams per liter per day for a period of atleast 30 days; the production bioreactor has a product residence time ofabout 1 to about 10 days; the production bioreactor has a dilution rateof about 1 to about 0.1 volume per day; the production bioreactor is fedwith a diluent solution; the diluent solution is water or saline.

In another aspect, the subject technology relates to a protein ofinterest produced by a method including: (a) culturing cells comprisinga gene that encodes the protein of interest in a culture bioreactor (N-1bioreactor); (b) inoculating a production bioreactor (N bioreactor) withcells obtained from step (a); and (c) culturing the cells in theproduction bioreactor under conditions that allow production of theprotein of interest. In one or more embodiments relating directly orindirectly to this aspect, the method further comprises step (d)harvesting the protein of interest from the production bioreactor; theculture bioreactor is a continuous perfusion culture bioreactor and thecontinuous production bioreactor is a continuously stirred tank reactor(CSTR) production bioreactor; the production bioreactor has no cellretention device; volume ratio of the culture bioreactor to theproduction bioreactor is about 1:1 to about 1:20; volume ratio of theculture bioreactor to the production bioreactor is about 1:1 to about1:5; volume ratio of the culture bioreactor to the production bioreactoris about 1:5; the inoculation in step (b) is by transferring cells fromthe culture bioreactor to the production bioreactor; the cell transferis by cell bleed in continuous or semi-continuous modes; step (a)optionally alternates between a first and second culture bioreactors toallow for renewal and continuous production of culture cells for use instep (b); the second culture bioreactor is a continuous perfusionculture bioreactor; the production bioreactor operates continuously fora period of greater than 3 weeks; the production bioreactor operatescontinuously for a period of greater than 4 weeks; the productionbioreactor operates continuously for a period of greater than 5 weeks;the production bioreactor operates continuously for a period of greaterthan 6 weeks; harvesting step (d) is continuous; the protein of interestis an antibody or a fusion protein; the cells are CHO cells, HEK-293cells, VERO cells, NSO cells, PER.C6 cells, Sp2/0 cells, BHK cells, MDCKcells, MDBK cells or COS cells; the production bioreactor is fed with adiluent solution; the diluent solution is water or saline.

Additional advantages of the subject technology will become readilyapparent to those skilled in this art from the following drawings andthe detailed description. The drawings and description are to beregarded as illustrative in nature, and not as restrictive.

BRIEF DESCRIPTION OF THE DRAWINGS

These and other features, aspects, and advantages of the claimedmethods, apparatuses, and/or systems are better understood when thefollowing detailed description is read with reference to theaccompanying drawings:

FIG. 1 is an example schematic illustrating the bioreactorconfiguration, according to an embodiment, described in detail in theexperimental section of this application. In this configuration an N-1bioreactor using perfusion with cell retention continuously suppliescells to a second bioreactor operating as a continuous-flow stirred-tankreactor (CSTR) or chemostat. The CSTR would also receive an additionalstream of concentrated nutrient feed. The CSTR would have no cellretention system and would be harvested continuously. In the experimentsperformed at the bench scale the cell retention device consisted of a0.2 micron micro-filtration hollow cartridge (General ElectricCFP-2-E-4X2MA) with a surface area of 850 cm². Cells were circulated at˜120 mL/minute through the lumen side of a hollow fiber cartridge by aWatson-Marlow peristaltic pump using 6.4 mm ID and 12.7 mm OD tubing.

FIG. 2A is a plot showing the viable cell density as measured by trypanblue exclusion as described in the experimental section of thisapplication. Open squares represent cell density in the N-1 perfusionreactor. Solid squares represent cell density in the CSTR or productionbioreactor. Dashed vertical lines represent step changes to a slightlydifferent perfusion medium composition flowing into the N-1 perfusionreactor as indicated in Table 2.

FIG. 2B is a plot showing the percent viable cells as measured by trypanblue exclusion. Open squares represent cell viability in the N-1perfusion reactor. Solid squares represent cell viability in the CSTR orproduction bioreactor.

FIG. 3 illustrates the perfusion rate and the cell bleed rate of the N-1bioreactor, and the dilution rate of the production (CSTR) bioreactor.The perfusion rate (open squares) is listed in reactor volumes (RV) perday and corresponds to the left hand y-axis. The perfusion rate is thevolume of medium that flowed into the N-1 perfusion bioreactor per day.The cell bleed rate (solid squares) is the volume of culture containingcells removed (continuously) from the N-1 perfusion bioreactor each dayand corresponds to the right hand y-axis. The dilution rate (solidtriangles) is the fraction of volume containing cells (the only streamleaving the CSTR) that is removed (continuously) from the production(CSTR) bioreactor each day and corresponds to the right hand y-axis.

FIG. 4 is a plot showing the residual glucose concentration in the N-1(open squares) and production (CSTR, solid squares) bioreactors.

FIG. 5 is a plot showing the lactate concentration in the N-1 (opensquares) and production (CSTR, solid squares) bioreactors.

FIG. 6 is a plot showing the antibody or product concentration in thebulk fluid of the N-1 bioreactor (solid squares), in the permeateleaving the N-1 bioreactor system (open squares), and in the bulk fluidand leaving the production bioreactor continuously (CSTR, solidtriangles).

FIG. 7 is a plot showing the volumetric productivity of the CSTRproduction bioreactor in grams product produced per reactor volume perday (solid triangles).

FIG. 8 is a plot showing the cell density of two N-1 perfusionbioreactors as a function of time. The target steady-state viable celldensity for the 1^(st) N-1 perfusion reactor was ˜40×10⁶ cells/mL, andfor the 2^(nd) N-1 perfusion reactor was ˜80×10⁶ cells/mL. As describedin more detail in Example 2, in this configuration, each of these N-1perfusion bioreactors independently feeds into a separate CSTRproduction bioreactor.

FIG. 9 is a plot showing the cell viability of two N-1 perfusionbioreactors as a function of time.

FIG. 10 is a plot showing the residual glucose as a function of time forthe two continuous perfusion bioreactors (open symbols) and for the twoproduction bioreactors operating as CSTR's (solid symbols) each linkedto one of the perfusion bioreactors.

FIG. 11 is a plot showing the lactate concentration as a function oftime for the N-1 continuous perfusion bioreactors (open symbols) and forthe production bioreactors operating as CSTR's (solid symbols).

FIG. 12 is a plot showing the perfusion rate of N-1 continuous perfusionbioreactors—perfusion rate is controlled by the cells using the HIPCOPtechnology as described in Example 2.

FIG. 13 is a plot showing the cell bleed rates of the two N-1 continuousperfusion bioreactors.

FIG. 14 is a plot showing the viable cell densities as a function oftime in the production bioreactors operating as continuous-flowstirred-tank reactors (CSTR's). Time intervals where dilution rates wereheld constant are denoted on the figure. The simulated volume ratio ofthe N-1 perfusion bioreactor to CSTR is also denoted on the figure.

FIG. 15 is a plot showing the dilution rate of each of the productionbioreactors operating as continuous-flow stirred-tank reactors (CSTR's)as a function of time.

FIG. 16 is a plot showing the cell viabilities as measured by trypanblue dye exclusion as a function of time in the production bioreactorsoperating as continuous-flow stirred-tank reactors (CSTR's).

FIG. 17 is a plot showing the product concentration or titer as afunction of time for the production bioreactors operating ascontinuous-flow stirred-tank reactors (CSTR's). Time intervals wheredilution rates were held constant are denoted on the figure.

FIG. 18 is a plot showing the osmolality as a function of time for theN-1 continuous perfusion bioreactors (open symbols) and for theproduction bioreactors operating as CSTR's (solid symbols). Denoted onthe figure is the time-point at which the perfusion medium was changedto contain the DL form of sodium lactate.

FIG. 19 is a plot showing the instantaneous volumetric productivity ofthe N-1 continuous perfusion bioreactors as considering operatingindependently. Calculated values assume that product is recovered fromboth the cell bleed the cell-free permeate leaving the cell retentionsystem.

FIG. 20 is a plot showing the concentration of product (antibody) in theN-1 perfusion reactors and in the cell-free permeate leaving the cellretention devices.

FIG. 21 is a plot showing the steady-state volumetric productivity ofthe linked N-1 continuous perfusion bioreactor to CSTR system, graphedas a function of production CSTR dilution rate. The volume ratio of N-1continuous perfusion bioreactor to production CSTR is denoted forseveral conditions. All points not specifically denoted used the 1:5volume ratio.

DETAILED DESCRIPTION OF THE SUBJECT TECHNOLOGY

In the following detailed description, numerous specific details are setforth to provide a full understanding of the subject technology. It willbe apparent, however, to one ordinarily skilled in the art that thesubject technology may be practiced without some of these specificdetails. In other instances, well-known structures and techniques havenot been shown in detail so as not to obscure the subject technology.

To facilitate an understanding of the present subject technology, anumber of terms and phrases are defined below:

Definitions:

The grammatical articles “one”, “a”, “an”, and “the”, as used herein,are intended to include “at least one” or “one or more”, unlessotherwise indicated. Thus, the articles are used herein to refer to oneor more than one (i.e., to at least one) of the grammatical objects ofthe article. By way of example, “a component” means one or morecomponents, and thus, possibly, more than one component is contemplatedand may be employed or used in an implementation of the describedembodiments.

The term “about” generally refers to a slight error in a measurement,often stated as a range of values that contain the true value within acertain confidence level (usually ±1σ for 68% C.I.). The term “about”may also be described as an integer and values of ±20% of the integer.

The terms “production bioreactor” or “N bioreactor” or “CSTR bioreactor”or “CSTR production bioreactor” as used interchangeably herein refer toa bioreactor (e.g., continuous-flow stirred-tank reactor (CSTR) thatdoes not utilize a cell retention system or device. The productionbioreactor is linked downstream to one or more culture bioreactor(s) andreceives inoculum from the culture bioreactor(s). The productionbioreactor is uniformly mixed and has a fluid in-flow that is equivalentto the fluid out-flow, therefore, maintaining a near constant volume.Such a production bioreactor will often achieve, although notnecessarily, a ‘chemically static’ or ‘steady-state’ environment whenoperating for sufficiently long periods of time, meaning that celldensities and other aspects of the culture (e.g., concentrations ofnutrients, etc.) will reach a stable (i.e. steady or static) state, andis therefore also commonly referred to as a ‘chemostat’. The fluidin-flow to and/or out-flow from the production bioreactor may beindependently continuous or semi-continuous. Such production bioreactoroperates continuously for a period or equal or greater than 3, 4, 5, or6 weeks.

In an embodiment, the CSTR production bioreactor of the subjecttechnology can continue to operate indefinitely or for as long as thecultured cells remain genetically stable. In another embodiment, theCSTR production bioreactor of the subject technology is linked to atleast two culture bioreactors which alternate in feeding fresh (i.e.,genetically stable) inoculum to the production bioreactor, and as suchthe production bioreactor continues to run for a long period of time,e.g., equal to or greater than 1, 2, 3, 4, 5, 6 months or for anindefinite period of time so long as the cultured cells that are beingfed to it remain genetically stable. Without a cell retention device,the fluid out-flow from the CSTR production bioreactor includes amixture of cells, cells debris and products. In an embodiment, thisfluid out-flow is directed to a cell separation device (e.g., an in-linecentrifuge) before being transferred to the next unit operation (e.g.,protein A column). The cells and cell debris separated in this fashionfrom the rest of the fluid out-flow will be discarded and not returnedto the production bioreactor.

The terms “cell retention system” or “cell retention device” as usedherein refer to a means of selectively retaining viable cells within thebioreactor such that the density of cells in the fluid leaving thebioreactor is lower than the density of cells in the fluid within thebioreactor. In this sense, a cell retention system or device isdifferent from a cell separation device mentioned above.

The terms “culture bioreactor” or “N-1 bioreactor” or “N-1 perfusionbioreactor” as used interchangeably herein refers to a perfusion cultureor perfusion bioreactor (e.g., continuous perfusion culture bioreactor)that is used for culturing the cells that will be used for inoculatingthe production bioreactor. Such culture bioreactor has a cell retentionsystem. There are many different forms of cell retention systems in usein industry. Some of these cell retention systems remove 100% of theviable cells from the liquid leaving the bioreactor system, however manyothers may remove only some variable fraction of the cells from theliquid leaving the bioreactor system. The liquid transfer to and fromthe culture bioreactor may be continuous or semi-continuous.Particularly, the cell bleed or transfer from the culture bioreactor tothe production bioreactor may be continuous or semi-continuous.

The term “semi-continuous,” in the context of liquid transfer to and/orfrom a bioreactor, as used herein means ‘periodic’ or refers to ascenario in which liquid (e.g., media alone and/or with cells, cellbleed) is added to and/or removed from the bioreactor once every howeverlong period of time. For example, once every 1, 2, 5, 10, 15, 30, 45 or60 minutes, or once every hour, or once every 2-3 hours, or once everyhowever long period of time from 1 minutes to 24 hours, a burst ofliquid is transferred from and/or to the bioreactor for a periodextending from few seconds (e.g., 1 sec., 2 sec., 5 sec. 10 sec., 20sec. or 60 sec.) to several minutes (e.g. 2 min. 5 min., 10 min., 25min., 50 min, 120 min. or 240 min.). The term “continuous” in thiscontext refers to a constant or non-periodic liquid transfer. In anembodiment, the rate of cell bleed/transfer from the culture bioreactorto the production bioreactor, either in continuous or semi-continuousmode, is less than the growth rate of the cells in the culturebioreactor. In another embodiment, the rate of cell bleed/transfer fromthe culture bioreactor to the production bioreactor is from about 0.1reactor volumes per day (RV/day) to about 1.3 RV/day. In anotherembodiment, the rate of cell bleed/transfer from the culture bioreactorto the production bioreactor, either in continuous or semi-continuousmode, is less than 1.3 RV/day, or is less than 1.0 RV/day, or is lessthan 0.8 RV/day, or is less than 0.6 RV/day, or is less than 0.4 RV/dayor is less than 0.2 RV/day to about 0.1 RV/day, or a varies within therange of about 0.1 to about 1.3 RV/day.

Perfusion culture of mammalian cells using a cell retention system canoffer a significant productivity advantage over batch, fed-batch, orchemostat/CSTR culture. Because cells are retained in the bioreactorsystem they can be perfused with large volumes of medium and can reachmuch higher cell densities without the concern for washout that wouldoccur in a CSTR at high dilution rates. Additionally, when a continuousperfusion bioreactor is linked to a continuous downstream purificationtrain, the size of the purification train can be reduced dramatically,which will simplify the overall operations by removing hold steps (andthe tanks required), and reducing sampling and analysis of in-processpools. However, perfusion culture of mammalian cells also suffers frommany disadvantages, particularly when implemented at large scale (>1,000L).

Typically most perfusion cultures are operated for long periods of time(>3 weeks). This requires that the cell line be particularly geneticallystable so that the cells continue to produce the protein of interest,and that the rate of epigenetic change is sufficiently slow. Theselong-duration perfusion cultures might be defined as ‘sustainable’perfusion cultures, in that the culture can be considered to be able torun as long as the genetic profile of the culture does not drift farenough to negatively affect the productivity or the product qualityprofile (specifications) of the protein of interest. In such asustainable perfusion culture, it is necessary to maintain high cellularviability, and this typically necessitates that the cells continue tomaintain a non-zero growth rate to continue to divide to make up forcells that die from apoptosis, or shear or oxidative or other stresses.In most cases a finite cell bleed rate (the removal of whole culturebroth containing cells) is required to maintain some cell growth and tocontinuously remove some non-viable biomass from the culture. This ‘cellbleed rate’ is usually referred to in terms of fraction of the culturevolume removed per day, or in terms of reciprocal days. Commonly cellbleed rates are in the 0.05-0.5/day range, with higher rates generallycreating conditions for higher cell viability, higher cell growth rates,and more sustainable perfusion systems, but simultaneously fosteringlower cell density and generally lower volumetric productivity perfusioncultures.

Additionally, depending upon the cell line, oftentimes CHO (Chinesehamster ovary) cell cultures will have higher specific rates ofproductivity (mass protein produced per cell per time) when the growthrate of the cells is low. This might be due to a cell devoting more ofits resources to continued cell division as opposed to producing theprotein of interest when growth rates are high. Fed-batch culture is oneof the most common modes of operation used for large-scale productionCHO cell cultures and typically will have a short period of cell growthfollowed by several days in a stationary phase in which most of theproduct is produced. It follows that there is an advantage to a culturesystem in which the phases of the culture are separated and theconditions for each phase are optimized; one phase with high cell growthrates to quickly obtain a high cell density, and a separate phase withnear zero cell growth rates, but much higher levels of cell specificproductivity.

Continuous perfusion systems can be difficult to implement at the largescale (>1,000 L) because cell retention systems do not always performwell at the large scale. Additionally, nearly any cell retention devicethat uses a membrane (as do many that are currently in use at largescale) will eventually plug with cell debris. This plugging is morelikely to occur with low cell viability cultures as the cells are morefragile, and the particulates that are formed are often similar in sizeto the pores of the membrane (0.2 micron). Therefore, there areadvantages to minimizing the size of the vessel in which the perfusionis performed, and keeping the cell viability high to minimize problemswith membrane fouling.

As mentioned above, cell particulates can slowly plug membrane-basedcell retention systems that utilize micro-filtration membranes onlong-term perfusion cultures. It is also possible that as cells lyse andhost-cell proteins and the product of interest interact and form highermolecular weight substances. These substances can also contribute to aphenomenon known as gel-layer formation, or product sieving. Under theseconditions, cell retention devices utilizing micro-filtration membranescan begin to act as ultra-filtration membranes and selectively retainhigh molecular weight proteins within the bioreactor system. Sinceimmunoglobulin molecules, one of the most common proteins produced incell cultures, have relatively high molecular weights (˜150,000 Da andabove), product sieving has emerged as a major issue with large scaleperfusion systems. Oftentimes a continuous perfusion system is linked toa continuous downstream purification system, so it is necessary for theprotein of interest to flow out of the bioreactor and for that proteinto ideally be of a consistent quantity and quality. The minimization ofprotein-of-interest retention is particularly important to minimize thesize of the downstream system, and is critical when the protein ofinterest is labile and likely to be degraded by excessive exposure tothe conditions inside the bioreactor.

For many of the reasons stated above, there are significant advantagesto continuously operating two bioreactors linked together in thefollowing manner and as shown in FIG. 1. The first bioreactor operatesas a continuous perfusion culture utilizing high perfusion rates whichallow for high cell densities. In this reactor high cell bleed rateswill also be used which will maintain high cell growth rates and alsotherefore high cell viabilities. The second bioreactor, the productionbioreactor, can advantageously be 4-5 times the volume of the firstbioreactor (typical volume ratios used in large scale facilities for theN-1 and production bioreactor) and will—according to a preferred mode ofthe subject technology—operate as a chemostat (or continuously stirredtank reactor [CSTR]) reactor with no cell retention system. The termsCSTR or chemostat will be used interchangeably to refer to theproduction bioreactor throughout this document as in most cases the CSTRwill quickly reach a steady-state or ‘chemically static’ condition inwhich most or all cell culture parameters will reach nearly constantvalues. In one embodiment, the volume ratio of the N-1 continuousperfusion culture bioreactor to the CSTR production bioreactor is about1:1 to about 1:20. In another embodiment, the volume ratio of the N-1continuous perfusion culture bioreactor to the CSTR productionbioreactor is about 1:1 to about 1:5. In another embodiment, the volumeratio of the N-1 continuous perfusion culture bioreactor to the CSTRproduction bioreactor is about 1:10. In another embodiment, the volumeratio of the N-1 continuous perfusion culture bioreactor to the CSTRproduction bioreactor is about 1:4.

In a first aspect, the subject technology relates to a method ofproducing a protein of interest, including: (a) culturing cellsincluding a gene that encodes the protein of interest in a culturebioreactor (N-1 bioreactor); (b) inoculating a production bioreactor (Nbioreactor) with cells obtained from step (a); and (c) culturing thecells in the production bioreactor under conditions that allowproduction of the protein of interest.

In a second aspect, the subject technology relates to a linked cultureprocess, including: (a) culturing cells including a gene that encodesthe protein of interest in a culture bioreactor (N-1 bioreactor); (b)inoculating a production bioreactor (N bioreactor) with cells obtainedfrom step (a); and (c) culturing the cells in the production bioreactorunder conditions that allow production of the protein of interest.

In a third aspect, the subject technology relates to a linked culturesystem, including: a culture bioreactor (N-1 bioreactor) for culturingcells including a gene that encodes the protein of interest; and aproduction bioreactor (N bioreactor) which is linked to the culturebioreactor and receives cells as inoculum from the culture bioreactor,wherein the production bioreactor operates under conditions that allowproduction of the protein of interest.

In a forth aspect, the subject technology relates to a protein ofinterest produced by a method including: (a) culturing cells including agene that encodes the protein of interest in a culture bioreactor (N-1bioreactor); (b) inoculating a production bioreactor (N bioreactor) withcells obtained from step (a); and (c) culturing the cells in theproduction bioreactor under conditions that allow production of theprotein of interest.

In one or more embodiments relating, directly or indirectly to any ofthe above aspects, the described method, process or system furtherincludes step (d) harvesting the protein of interest from the productionbioreactor; the culture bioreactor is a continuous perfusion culturebioreactor and the production bioreactor is a continuously stirred tankreactor (CSTR) production bioreactor; the production bioreactor has nocell retention device; volume ratio of the culture bioreactor to theproduction bioreactor is about 1:1 to about 1:20; volume ratio of theculture bioreactor to the production bioreactor is about 1:1 to about1:5; volume ratio of the culture bioreactor to the production bioreactoris about 1:4; the inoculation in step (b) is by transferring cells fromthe culture bioreactor to the production bioreactor; the cell transferis by cell bleed in continuous or semi-continuous modes; the celltransfer is in semi-continuous mode including the cell transfer once atevery period of time from 2 minutes to 24 hours or at any intervaltherebetween; step (a) optionally alternates between a first and secondculture bioreactors to allow for renewal and continuous production ofculture cells for use in step (b); the second culture bioreactor is acontinuous perfusion culture bioreactor; the production bioreactoroperates continuously for a period of greater than 3 weeks; theproduction bioreactor operates continuously for a period of greater than4 weeks; the production bioreactor operates continuously for a period ofgreater than 5 weeks; the production bioreactor operates continuouslyfor a period of greater than 6 weeks; the harvesting step (d) iscontinuous; the cells are for example CHO cells, HEK-293 cells, VEROcells, NSO cells, PER.C6 cells, Sp2/0 cells, BHK cells, MDCK cells, MDBKcells or COS cells or any cell genetically derived therefrom like forexample derivatives with specific metabolic conditions and/or selectionsystems like the glutamine selection, which genetically derived cellsand/or metabolic conditions and/or selection systems are in principleknown to a person skilled in the art; the production bioreactor has avolumetric productivity of at least 0.6 grams per liter per day for aperiod of at least 14 days; the production bioreactor has a volumetricproductivity of at least 0.6 grams per liter per day for a period of atleast 20 days; the production bioreactor has a volumetric productivityof at least 0.6 grams per liter per day for a period of at least 30days; the production bioreactor has a product residence time of about 1to about 10 days; the production bioreactor has a dilution rate of about1 to about 0.1 volume per day.

Non-limiting examples of mammalian cells, which can be used for thepresent invention are summarized in Table 1.

TABLE 1 Suitable exemplary mammalian production cell lines CELL LINEREFERENCE NUMBER NSO ECACC No. 85110503 Sp2/0-Ag14 ATCC CRL-1581 BHK21ATCC CCL-10 BHK TK⁻ ECACC No. 85011423 HaK ATCC CCL-15 2254-62.2 (BHK-21derivative) ATCC CRL-8544 CHO ECACC No. 8505302 CHO wild type ECACC00102307 CHO-DUKX (═CHO duk⁻, ATCC CRL-9096 CHO/dhfr⁻) CHO-DUKX B11 ATCCCRL-9010 CHO-DG44 Urlaub et al., Cell 33 (2), 405-412, 1983; LifeTechnologies A1097101 CHO Pro-5 ATCC CRL-1781 CHO-S Life TechnologiesA1136401; CHO-S is derived from CHO variant Tobey et al. 1962 Lec13Stanley P. et al, Ann. Rev. Genetics 18, 525-552, 1984 V79 ATCC CCC-93HEK 293 ATCC CRL-1573 COS-7 ATCC CRL-1651 HuNS1 ATCC CRL-8644 Per. C6Fallaux, F. J. et al, Human Gene Therapy 9 (13), 1909-1917, 1998 CHO-K1ATCC CCL-61, ECACC 85051005 CHO-K1/SF ECACC 93061607 CHO-K1 GS glutaminesynthetase (GS) deficient cells derived from CHO-K1 CHOZN GS SAFC ECACC85051005, cells derived from CHO-K1

Said production cells are cultivated preferentially under conditionsthat allow the cells to proliferate. Furthermore, said production cellsare cultivated preferentially under conditions, which are favorable forthe expression of the desired gene(s) and/or the protein of interest.The protein of interest is than isolated from the cells and/or the cellculture supernatant. Preferably the protein of interest is recoveredfrom the culture medium as a secreted polypeptide, or it can berecovered from host cell lysates if expressed without a secretorysignal.

Culture from the N-1 bioreactor will enter the production bioreactor ata continuous and fixed flowrate which will be equivalent to the cellbleed rate of the N-1 bioreactor. Additionally, the productionbioreactor will have a continuous feed of nutrient media as the cells inthat bioreactor will continue to metabolize, produce product, andpotentially undergo some limited cell division. The working volume ofboth bioreactors would likely be held constant, so the effectivedilution rate of the production bioreactor would be determined by theflows of the N-1 continuous cell bleed and the continuous feed ofnutrient media and any pH-controlling titrant added directly to theproduction reactor.

Such a bioreactor system might mitigate many of the problems of currentcontinuous perfusion bioreactors. The perfusion bioreactor and itsassociated cell retention system will likely be significantly simpler tooperate as they will operate at approximately one-fifth scale of theproduction bioreactor. If gel-layer/product sieving occurs in the cellretention device, this will only increase the overall productivity ofthe system as now the protein of interest generated in the N-1bioreactor will flow into the production bioreactor and eventually tothe downstream purification process rather than to drain with cell-freepermeate as shown in FIG. 1.

In situations where the cell line is not sufficiently genetically stablefor long term operation, a second N-1 bioreactor will be used that wouldstart periodically with an inoculum of fresh cells expanded from frozenvials and once that bioreactor has come up to the proper cell densitythe second N-1 bioreactor could take the place of the first N-1bioreactor while the first N-1 bioreactor is taken down for cleaning andre-sterilization. The cell density in the N-1 culture bioreactor(s) mayvary according to the size of the production bioreactor is it linked to.In some embodiments, the cell density in the culture bioreactor is about10×10⁶ per liter or more, or is about 20×10⁶ per liter or more, or isabout 40×10⁶ per liter or more, or is between 10×10⁶ and 200×10⁶ perliter, or is between 40×10⁶ to 120×10⁶ per liter or is at a specificdensity within any of these ranges. Since the cell division in theproduction bioreactor is likely to be low, this would allow for asemi-continuous renewal of the cells producing the protein with agenetically younger cell population. The production bioreactor mightthen be able to operate continuously with minimal interruption and athigh cell density and high productivity, producing a feed stream todownstream purification operations with consistent quality parametersfor months at a time. Therefore, in one embodiment, the linked cellculture system of the subject technology includes an N-1 continuousperfusion culture bioreactor (N-1 bioreactor) in which cells thatinclude a gene that encodes the protein of interest are cultured beforethey are transferred to a CSTR production bioreactor without a cellretention device. In another embodiment, the linked cell culture systemof the subject technology includes two N-1 culture bioreactors thatoperate in alternative to produce cultured cells for transfer to a CSTRproduction bioreactor. In another embodiment, the N-1 continuousperfusion culture bioreactor(s) of the subject technology operate in thefollowing modes: hi-end pH-control of glucose (HIPDOG), described inBiotechnol Bioeng. 2011 June; 108(6):1328-37, or high-end pH-control ofperfusion rate (HIPCOP) described in co-pending application WO2016/196261, entitled “Cell-Controlled Perfusion in Continuous Culture”for a description of this method.

Cell lines with higher specific productivities at low growth rates, orcells that require different cell culture environmental conditions to behighly productive or produce product with particular product qualitycharacteristics, should particularly benefit from such a bioreactorconfiguration as it is likely that the growth rate in the productionbioreactor will be quite low and more similar to the conditions that aretypically achieved in the later stages of a fed-batch bioreactor. Thisalso might mean that the product quality of the protein of interestproduced in such a linked continuous N-1 to CSTR production bioreactormight be more similar to the product quality of fed-batch producedprotein. Therefore, in an embodiment, the CSTR production bioreactoroperates under conditions that promote highest cell productivity whichmight in some cases also be conditions which slow cell growth. Toachieve high productivity conditions, known chemicals that improve percell productivity but slow or stop growth can be added to the CSTRproduction bioreactor. Alternatively, the CSTR bioreactor may be made tooperate, for example, under low or high pH, or low temperature (withadditions of Cu, low Ca, additions of galactose, etc.) that arebeneficial for getting cells to produce protein with particular qualityattributes, but those same conditions do not allow for high rates ofcell growth.

As mentioned above, the cell specific productivity might in some casesbe inversely related to the growth rate of the cells in the production(CSTR) bioreactor. A CSTR operating at a high dilution rate shouldresult in higher growth rates as inhibitory metabolites would be moreeffectively flushed out of the bioreactor. The dilution rate of the CSTRwill be dependent upon the cell bleed rate from the N-1 continuousperfusion bioreactor, and the rate of feed and any titrant used for pHcontrol directly entering the CSTR. Manipulating the concentration ofthe feed medium to the CSTR would allow for a balance of optimal growth,nutrient availability, and inhibitor metabolite flushing from thesystem. Multiple factors, many of which are interrelated mightcontribute to the maximum productivity of the combined N-1 continuousperfusion and CSTR system, including the cell density and growth ratesof the cells in both bioreactors, the perfusion rates in the N-1, andthe dilution rates of the CSTR. It is advantageous to manipulate thedilution rate of the CSTR by using a single concentrated feed media, andto dilute this media with water, or water mixed with a saturated salinesolution as needed. Control of the dilution rate with variableconcentration feed media might also allow for control over the residencetime of the recombinant protein produced from the system. High dilutionrates would result in low product residence times and might beadvantageous for highly labile proteins.

Many in industry are currently exploring the use of cellretention/perfusion bioreactors as the production system for mammaliancell cultures. Such systems increase volumetric productivity and makepossible continuous downstream purification trains. Since as with acontinuous perfusion bioreactor as the production bioreactor, the linkedN-1 perfusion bioreactor to chemostat or CSTR production bioreactordescribed in the present work would continue to generate a constantstream of product to the downstream operation, the advantages of asmaller and continuous downstream operation could continue to berealized.

In addition, the subject technology relates to a protein of interestproduced by a method as described above. Such protein of interestincludes, but is not limited to an antibody or a fusion protein, such asa Fc-fusion proteins. Others can be for example enzymes, cytokines,lymphokines, adhesion molecules, receptors and derivatives or fragmentsthereof, and any other polypeptides and scaffolds that can serve asagonists or antagonists and/or have therapeutic or diagnostic use. Otherrecombinant proteins of interest are for example, but not limited toinsulin, insulin-like growth factor, hGH, tPA, cytokines, such asinterleukins (IL).

A preferred recombinant secreted therapeutic protein is an antibody or afragment or derivative thereof. Thus, the invention can beadvantageously used for production of antibodies such as monoclonalantibodies, multispecific antibodies, or fragments thereof, preferablyof monoclonal antibodies, bi-specific antibodies or fragments thereof.Antibody fragments include e.g. “Fab fragments” (Fragmentantigen-binding=Fab). Fab fragments consist of the variable regions ofboth chains, which are held together by the adjacent constant region.These may be formed by protease digestion, e.g. with papain, fromconventional antibodies, but similar Fab fragments may also be producedby genetic engineering. Further antibody fragments include F(ab')2fragments, which may be prepared by proteolytic cleavage with pepsin.

Using genetic engineering methods it is possible to produce shortenedantibody fragments which consist only of the variable regions of theheavy (VH) and of the light chain (VL). These are referred to as Fvfragments (Fragment variable=fragment of the variable part). Since theseFv-fragments lack the covalent bonding of the two chains by thecysteines of the constant chains, the Fv fragments are often stabilized.It is advantageous to link the variable regions of the heavy and of thelight chain by a short peptide fragment, e.g. of 10 to 30 amino acids,preferably 15 amino acids. In this way a single peptide strand isobtained consisting of VH and VL, linked by a peptide linker. Anantibody protein of this kind is known as a single-chain-Fv (scFv).Examples of scFv-antibody proteins are known to the person skilled inthe art. Preferred secreted recombinant therapeutic antibodies accordingto the invention are bispecific antibodies. Bispecific antibodiestypically combine antigen-binding specificities for target cells (e.g.,malignant B cells) and effector cells (e.g., T cells, NK cells ormacrophages) in one molecule. Examplary bispecific antibodies, withoutbeing limited thereto are diabodies, BiTE (Bi-specific T-cell Engager)formats and DART (Dual-Affinity Re-Targeting) formats. Also anticipatedin the context of the present invention are minibodies. By minibody, theskilled person means a bivalent, homodimeric scFv derivative.

The recombinant secreted therapeutic protein, especially the antibody,antibody fragment or Fc-fusion protein is preferably recovered/isolatedfrom the culture medium as a secreted polypeptide. It is necessary topurify the recombinant secreted therapeutic protein from otherrecombinant proteins and host cell proteins to obtain substantiallyhomogenous preparations of the recombinant secreted therapeutic protein.As a first step, cells and/or particulate cell debris are removed fromthe culture medium or lysate. Further, the recombinant secretedtherapeutic protein is purified from contaminant soluble proteins,polypeptides and nucleic acids, for example, by fractionation onimmunoaffinity or ion-exchange columns, ethanol precipitation, reversephase HPLC, Sephadex chromatography, and chromatography on silica or ona cation exchange resin such as DEAE. Methods for purifying aheterologous protein expressed by host cells are known in the art.

EXAMPLES Example 1 Application of the Continuous N-1 PerfusionBioreactor Feeding a CSTR Production Bioreactor (at the 1-2 liter scale)in Producing a Humanized IgG with CHO Cells

FIG. 1 shows the diagram of the two-stage linked bioreactor system usedin the experiment. For the purposes of the experiment the N-1 perfusionbioreactor working volume including the perfusion loop (cell-retentionsystem) was 1.25 liters and the working volume of the productionbioreactor was 1.0 liters. At the full scale in an industrial setting,it is contemplated that the N-1 bioreactor would be approximately onefifth the volume of the production (CSTR or chemostat) bioreactor. Forthis reason the experimental bioreactors were operated to simulate sucha volume ratio. This means that the majority of cell bleed (-81% of thetotal cell bleed) from the N-1 perfusion bioreactor was sent to drain.Only the appropriate volume (˜19% of the total) of the cell bleed waspumped into the production bioreactor. The only exception to this ruleoccurred in the first 3 days of operation of the production (CSTR)bioreactor in which all of the cell bleed from the N-1 was added to theproduction bioreactor. This allowed the production bioreactor to come upto a high density slightly faster than would have otherwise occurred.

FIG. 2a initially shows only the N-1 perfusion bioreactor cell density.That density quickly increased to the target density of approximately40×10⁶ viable cells/mL, after which the cell bleed was initiated andmanually adjusted over the course of the culture to achieve a constantviable cell density in the N-1 bioreactor of approximately 40×10⁶ viablecells/mL. As shown in FIG. 3, for the majority of the experiment thecell bleed rate of the N-1 perfusion bioreactor was approximately0.4/days, or 0.4 reactor volumes removed per day.

The N-1 perfusion bioreactor used a feedback mechanism based on pH whichallows the culture to determine its own rate of perfusion. The perfusionmedium contained a small amount of sodium-L-lactate (listed in Table 2).

TABLE 2 lists details of the compositions of the perfusion media thatwas used for the N-1 bioreactor.

TABLE 2 The compositions of the perfusion media in the N-1 bioractor.Total Osmotic Used from AA Glucose Sodium-L- Strength Medium days (mM)(g/L) Lactate (g/L) (mOsm) Rich N-1 1-15 and 70 5.3 1.8 335 perfusion29-35 Lean N-1 15-29 60 4.2 1.7 319 perfusion

The culture ‘signals’ a lack of available glucose by beginning to takeup lactic acid from the bulk culture. The removal of lactic acid by thecells from the bulk culture fluid results in a rise in the pH of thebioreactor which in turn activates a pump which delivers perfusionmedium containing glucose to the N-1 perfusion bioreactor, therebyincreasing the perfusion rate of the culture (a level controller removescell-free medium [permeate] from the culture through the shell side ofthe hollow fiber cell retention device to maintain a constant bioreactorvolume). When the cells in the culture begin to take up the excessglucose that is being delivered, some fraction of it is converted tolactic acid by the cells which then suppresses the bulk culture pH as itis secreted from the cells, which in turn causes the pH controller todeactivate the perfusion medium addition pump (also in turn deactivatingthe permeate pump). This control scheme occurs again and again as acycle every few minutes during the course of an experiment. Seeco-pending application No. WO 2016/196261, entitled “Cell-ControlledPerfusion in Continuous Culture” for a description of this method.Briefly, in the method described in co-pending application No. WO2016/196261, the pH trigger for turning on and off the perfusion pumps,fresh medium and permeate pumps, is set at a predetermined value. Thispredetermined value is about pH 7 (e.g., between 6.8-7.4). The mediumcomprises glucose, L-lactate and a specified ratio of amino acid toglucose. The described method also comprises adding L-lactate to a freshperfusion medium used in the continuous culture process. The L-lactatepresent in the perfusion medium is in an amount of about 0.1 g/L to 7.0g/L. Alternatively or in addition, the L-lactate present in theperfusion medium is in an amount of about 1 to 4 g/L, about 1 to 3 g/L,or about 1 to 2.5 g/L. In a preferred embodiment, the L-lactate issodium L-lactate or potassium L-lactate. In another embodiment, thelactate used in the perfusion medium may be a D/L mixture of sodium orpotassium lactate. In cases where a D/L mixture of lactate is used,enough mixture is added to the perfusion medium such that the amount ofL-lactate would be identical to that used when only L-lactate was used.In an embodiment, additional sodium bicarbonate is added to theperfusion medium instead of, or in addition to, L-lactate. A typicalcell culture medium contains about 1.0 to 2.5 grams/L of sodiumbicarbonate. In an embodiment, the perfusion medium for the N-1perfusion bioreactor of the subject technology has an additional 1 to 3grams/L of sodium bicarbonate to achieve a total concentration of about2 to 5.5 grams/L of sodium bicarbonate. The perfusion medium in thismethod requires at least glucose, L-lactate (or alternatively or inaddition, sodium bicarbonate) and amino acids. The concentration ofglucose is about 0.5 to about 40 g/L. The concentration of L-lactate isabout 0.1 to about 7.0 g/L. The concentration of sodium bicarbonate isabout 1 to 5.5 grams/L. The ratio of moles of glucose to amino acids isbetween about 0.25 and 1.0. In HIPCOP model, as the cell densityincreases (FIG. 2a ) the cells effectively ‘request’ additionalperfusion medium on a more and more frequent basis resulting in a rampup of perfusion rate (FIG. 3).

A similar control scheme for the addition of glucose to fed-batchcultures has been described in the literature (Biotechnol Bioeng. 2011June; 108(6):1328-37). The N-1 perfusion culture uses this technology ofhigh-end pH-control of perfusion rate (HIPCOP) for the entire durationof the experiment. When necessary or useful, the CSTR or productionbioreactor also uses a similar technique of glucose limitation tocontrol lactic acid formation. In the CSTR when the pH rises as a resultof lactic acid uptake by the cells, a pump is triggered to add aconcentrated nutrient feed containing glucose, or a feed containing pureglucose dissolved in water. Table 3 lists the feed media or pure glucosesolution that was used as part of this hi-end pH-control of glucose(HIPDOG) strategy.

TABLE 3 lists details of the compositions of the concentrated feed mediathat was delivered to the CSTR (chemostat) on-demand via the high-endpH-control system (HIPDOG) when that system was in operation from days12-19, and again from days 29-33.

TABLE 3 The compositions of the concentrated feed media. Total OsmoticUsed from AA Glucose Sodium-L- Strength Medium days (mM) (g/L) Lactate(g/L) (mOsm) Lean Feed to 12-16 380 80 0 ~1105 production CSTR Rich Feedwith 16-17, 18-19 600 120 0 ~1702 high glucose to production CSTR RichFeed with 17-18 600 80 0 ~1480 low glucose to production CSTRConcentrated 29-33 0 300 0 ~1666 glucose feed to production CSTR

The HIPDOG strategy was used to control lactic acid formation for theCSTR only from days 12-19, and from days 29-33. Additional fixed ratefeeding of concentrated nutrient solutions occurred to the CSTR fromdays 19-35. The rate of addition of these feeds was determined byoccasionally analyzing off-line samples for residual amino acid levelsin the CSTR and changing the feed rate to the CSTR if excess (above 30millimolar) total amino acids or depleted (below 30 millimolar) totalamino acids were detected. Using the feeding schemes described in thisparagraph we believe that the cells in the CSTR were not likely overlylimited with respect to glucose availability, nor were they limited foravailability of any particular amino acid. When steady-state, or nearsteady-state conditions are reached in the N-1 and CSTR combinedreactors between about day 25-28 and also day 32-35, from FIG. 3 we cansee that the cell bleed of the N-1 bioreactor is approximately 0.43/dayand the dilution rate of the production CSTR reactor is approximately0.13-0.15/day. In some embodiments, the dilution rate of the productionbioreactor is about 0.05-0.4 or about 0.1-0.3 or about 0.15-0.2/day. Dueto the ⅕th volume ratio that was simulated with these two linkedbioreactors it can be calculated that the cell bleed from the N-1perfusion reactor is contributing approximately 53-62% of the volume ofliquid entering the CSTR, and the remainder consists of concentratedmedia or glucose feeds.

The line from the N-1 perfusion bioreactor to the production bioreactor(CSTR or chemostat) was first connected and opened on about day 12 whichallowed cells to begin to flow into the CSTR. In an industrial settingthere would be no strong driver to delay the start of the CSTR.Presumably it could be started at the first time point that cell bleedis removed from the N-1 perfusion bioreactor. In the current experimentthe CSTR was started after it was believed that the N-1 had reached anear steady-state condition.

From day 12 to 15 (the first 3 days of the CSTR operation) all of thecell bleed from the N-1 was pumped into the production (CSTR)bioreactor. This would likely be the normal start up method atindustrial scale in which the N-1 bioreactor is approximately ⅕th thevolume of the CSTR. In the small scale experimental system however thevolume of the bioreactors are nearly identical and this means that amuch larger number of cells were added to the CSTR system in the first 3days of its operation than could likely occur at the large scale. Thiswas performed at the small scale to accelerate reaching a high densityin the CSTR, but should have negligible effect on the finalsteady-states reached in the combined bioreactor system. On day 15 theflow from the N-1 perfusion to the production CSTR reactor was adjustedto accurately simulate a 1:5 volume ratio.

Steady-state with respect to most parameters, cell density (FIG. 2a ),cell viability (FIG. 2b ), bioreactor volume flows (FIG. 3), metaboliteconcentrations (FIGS. 4 and 5), and product concentrations andvolumetric productivities (FIGS. 6 and 7) was first reached on about day23. At this point the volumetric productivity of the CSTR wasapproximately 0.9 grams/liter reactor volume/day. This volumetricproductivity was maintained or slightly exceeded for the next 12 days(FIG. 7). While some product sieving, or selective retention of antibodyproduct occurs in the N-1 bioreactor (solid and open squares of FIG. 6)across the hollow-fiber cell retention device, a simple mass balancecalculation considering the volume of liquid per day entering the CSTRfrom the N-1 bioreactor (˜65% of the total volume entering the CSTR) andits product concentration (0.4-0.6 grams/L), and the volume per day andconcentration of product (6.4-7.4 grams/L) leaving the CSTR on day 23shows that the volumetric productivity calculated for the CSTR isprimarily (˜93-98%) due to generation of product in the CSTR. TABLE 4lists details of the compositions of the concentrated feed media thatwas delivered to the production (CSTR) reactor at various fixed ratesvia manual manipulation from days 19-35.

TABLE 4 The compositions of the concentrated feed media in theproduction reactor. Total Osmotic Used from AA Glucose Sodium-L-Strength Medium days (mM) (g/L) Lactate (g/L) (mOsm) Rich Feed with19-27 600 50 0 ~1313 very low glucose to production CSTR Lean Feed to27-29 380 80 0 ~1105 production CSTR Lean Feed with 29-35 380 50 0 ~939low glucose to production CSTR

It is of value to compare the steady-state volumetric productivity ofthe combined bioreactor system (-0.9 grams/liter/day) on day 23 to thatof the N-1 perfusion bioreactor operating independently as a productionbioreactor. Using the product concentrations and volumes leaving the N-1perfusion bioreactor, and assuming that a downstream operation cancapture material from both streams (cell-free permeate andcell-containing cell bleed) the volumetric productivity of the N-1perfusion bioreactor on day 23 is approximately 0.46 grams/liter/day,roughly half of the volumetric productivity of the combined bioreactorsystem (0.9 grams/liter/day). Additionally, the medium consumption rateof the N-1 perfusion bioreactor operating independently on day 23 isapproximately 1.95 reactor volumes/day. The combined N-1/CSTR systemmedium consumption rate (including both the N-1 perfusion medium and theconcentrated nutrient feeds to the CSTR system) is only 0.44 reactorvolumes/day (based on the volume of the CSTR), less than one quarter ofthe medium consumption rate of the N-1 perfusion bioreactor operatingindependently as a production bioreactor.

It should be noted that while in this particular experiment the high-endpH-control of perfusion rate (HIPCOP) was used for the N-1 perfusionbioreactor, and the hi-end pH-control of glucose (HIPDOG) strategy wasintermittently used in the CSTR production bioreactor, these controlmodes are to be regarded as optional in carrying out the subjecttechnology. Alternative or additional control mechanisms may beapplicable by taking into consideration the metabolic conditions neededfor each cell type to be cultured by the methods and system of thesubject technology. For example, various CHO cell lines and othermammalian cell expression systems have distinctive cellular metabolismsthat might make the use of glucose limiting techniques optional for thelinked bioreactor described herein.

Example 2 IgG (Immunoglobulin G) Production Utilizing Cell Line B—aGlutamine-Synthetase Expression System CHO (Chinese Hamster Ovary) CellLine in Making an IgG Protein

The experimental design of the two-stage linked bioreactor system usedin this example is identical to that used in example 1. For the purposesof the experiment the N-1 perfusion bioreactor working volume includingthe perfusion loop (cell-retention system) was 1.36 liters and theworking volume of the production bioreactor (CSTR) was 1.1 liters. In anindustrial setting, the N-1 bioreactor is contemplated to beapproximately one fifth the volume of the production (CSTR or chemostat)bioreactor. For this reason (in a similar manner to example 1), theexperimental bioreactors were operated to simulate such a volume ratioat the beginning of the experiment. Later in the experiment (day 32) theexperimental bioreactor configurations were changed in a way such thatthey simulate a volume ratio of 1:10, and changed yet again on day 44 tosimulate a volume ratio of 1:20. In each of these cases, solely for thepurpose of simulating the theoretical bioreactor volume differences, alarge fraction of the cell bleed from the N-1 continuous perfusionbioreactor was sent to drain instead of going into the productionbioreactor. As the volume ratio simulation was increased from 1:5 to1:10 and then later to 1:20, a larger fraction of the volume of cellscoming from the N-1 perfusion bioreactor was sent to drain.

Various process control parameters and set-points (pH, dissolved oxygen,temperature, etc.) are listed in table 5 for the bioreactors used inthese experiments.

TABLE 5 Operating parameters and control set-points used in the N-lcontinuous perfusion and CSTR production bioreactors. N-1 PerfusionProduction CSTR Parameter Bioreactors Bioreactors Inoculation density(viable cells/mL) 1.2 × 10⁶ 0    Temperature (° C.) 36.5 Controlled withelectric heating blankets High-end pH Set-Point 7.125 7.125 (Controlledusing HIPCOP (Controlled using HIPDOG technology [see text]) technologyprior to day 22 or 26 [see text]) Low-end pH Set-Point 7.075 7.075(using 1M sodium (used only during first three carbonate/potassiumcarbonate days of culture) [molar ratio of 0.94 sodium:0.06 potassium])Dissolved Oxygen Set-Point 40 ± 10% (percent of air saturation) OxygenDelivery Method Pure oxygen delivered through 15 micron sintered-steelsparge tip Carbon Dioxide Removal Method Air delivered throughdrilled-hole sparger (large bubbles) 7 × 1 mm holes Flow rate between2.5 to 4X of the 15 micron sparge oxygen Working Volume (liters) 1.36 1.10  Agitation Single Rushton impeller (6 cm diameter) operating at 275RPM ~90 W/m³ power/unit volume

As mentioned in Table 5 a dual sparging strategy was used in which themajority of oxygen to the culture was delivered through a 15 micronsintered steel sparger using pure oxygen, and the majority of removal ofcarbon dioxide was accomplished by sparging atmospheric air through adrilled hole (7×1 mm holes) sparger producing large bubbles. The spargerate of the air through the drilled hole sparger was varied between 2.5to 4 times the volume of oxygen used in the 15 micron sparger. Thisstrategy was sufficient to control dissolved oxygen at the 40% of airsaturation set point, and to keep dissolved carbon dioxide levelsbetween 5-13% for the first 13 days of the cultures, and between 4-8%from day 13 on (when the majority of key data was collected).

For this example two N-1 perfusion bioreactors were operated and each ofthese reactors independently supplied a continuous cell source to oneproduction bioreactor operating as a continuous-flow stirred tankreactor (CSTR). The target steady-state viable cell density for the1^(st) N-1 perfusion reactor was ˜40×10⁶ cells/mL, and for the 2^(nd)N-1 perfusion reactor was ˜80×10⁶ cells/mL. The N-1 perfusionbioreactors were started eight days before the first cells weretransferred to the production bioreactors. The production bioreactors(the CSTR's) were started at full volume, meaning that the bioreactorswere completely full of medium when the first cell bleeds from the N-1perfusion bioreactors started to flow into the production bioreactors.The production bioreactors were started at full volume partly becausethe experimental plan for the linked bioreactor system experiment wasdesigned to explore several different effective dilution rates in theproduction bioreactors (CSTR's), and to attempt to collect steady-statedata for each of those conditions. This type of experiment (a dilutionrate study) is generally facilitated by exploring high dilution ratesfirst before moving to lower dilution rates that typically requirelonger time periods to reach steady-state conditions. Also, low dilutionrates are likely to result in lower viability cultures and could resultin long lag times for cells to react when later moving to higherdilution rates with more favorable culture conditions. For this reason,high dilution rates in the production bioreactors were investigatedfirst.

Starting the production bioreactor at full volume (prior to the additionof cells from the N-1 perfusion bioreactor) is probably not the optimalway to initiate such a culture. There may be a benefit from a medium andfacility utilization standpoint to start the bioreactor at partialvolume so that when the first material is removed from the productionbioreactor (when the bioreactor is full) the cell density, concentrationof product of interest, and potentially product quality of the proteinis already near to the final steady-state value. This optimal startingvolume (which depends upon the growth and product production kinetics ofany particular cell line) can be determined by routine experimentationand computer modeling simulations. The optimal starting volume of theproduction bioreactor also depends upon the optimal (highestproductivity and practical operating conditions) steady-state conditionsof the N-1 perfusion and CSTR combined system.

The two N-1 continuous perfusion bioreactors were inoculated at ˜1.2×10⁶viable cells/mL from cells maintained in Erlenmeyer shake flasks. Thebioreactors initially operated in a batch mode (no perfusion). As thecell density increased (FIG. 8) the glucose concentration fell fromaround 4 grams/L to below 0.4 grams/L (FIG. 10) on day 3. Simultaneouslylactate initially accumulated to slightly over 2 grams/L (FIG. 11). Oncethe glucose level fell to a sufficiently low concentration, the cellsbegan to take up lactic acid from the culture (day 3), raising the pH ofthe bulk fluid and triggering the addition of perfusion medium (and thesimultaneous removal of cell-free permeate through the cell retentionsystem to maintain the bioreactor working volume constant). See theco-pending application No. 62/246,774, entitled “Cell-ControlledPerfusion in Continuous Culture” for a description of this method. Thismethod will be referred to as HIgh-end pH Control Of Perfusion (HIPCOP).The cells continued to control and ramp up their own perfusion rate overthe next several days (FIG. 12). After approximately day 8 the perfusionrate stabilized for the two bioreactors and remained relatively stablefor the duration of the experiment. During the entire experiment theperfusion rate was controlled by the HIPCOP technology. The averageperfusion rate for the target 40×10⁶ cells/ml N-1 perfusion bioreactorfrom day 10 to day 51 can be calculated to be 1.76 reactor volumes perday (this is defined as the total volume of medium pumped into thebioreactor). The average perfusion rate for the target 80×10⁶ cells/mlN-1 perfusion bioreactor from day 10 to day 40 can be calculated to be3.60 reactor volumes per day. Beyond day 5 the residual glucose level inboth N-1 bioreactors was very near zero for the entire experiment as istypical of bioreactors operating with the HIPCOP technology (FIG. 10).

On day 6 the direct removal of cell containing culture (the cell bleed)from the N-1 perfusion reactors was initiated. Most long-duration orsustainable perfusion bioreactor operations require some amount of cellbleed to help maintain cell viability in the bioreactor and keep theoverall levels of inert biomass from becoming problematic. The cellviability for the N-1 perfusion bioreactors is shown in FIG. 9. Theviability for both cultures was high, above 90% for the entire length ofthe experiment. The viability was slightly lower in the higher densityculture. This might have been due to slightly higher shear forces thatwere likely experienced due to the almost twice as high gas sparginglevels that were necessary to maintain dissolved oxygen and carbondioxide in appropriate ranges in the higher density culture.

From days 6 to 8 the cell bleed was directed to waste as it was thoughtthat for the purposes of the experiment it might be easier to initiatethe production reactors (CSTR's) when the cell bleed rate and celldensities in the N-1 perfusion cultures had stabilized. The cell bleedrates were adjusted daily in an attempt to keep the cell densities ofthe N-1 perfusion reactors (FIG. 8) as close as possible to their targetdensities of 40 and 80×10⁶ cells/mL. The cell bleed rates are indicatedin FIG. 13. After day 8 the manual adjustments were minor and theaverage cell bleed rate for the target 40×10⁶ cells/ml N-1 perfusionbioreactor from day 10 to day 51 can be calculated to be 0.72 reactorvolumes per day. The average cell bleed rate for the target 80×10⁶cells/ml N-1 perfusion bioreactor from day 10 to day 40 can becalculated to be 0.74 reactor volumes per day.

On day 8 the cell bleeds from the N-1 perfusion reactors were directedinto the two independently operating production bioreactors (CSTR's). Asmentioned earlier in the text, the correct volume of cell bleed wasdirected into the production bioreactors such that a 1:5 volume ratio ofN-1 to production bioreactor was simulated. In an industrial applicationin which the trajectories of cell densities and cell bleed rates, andthe final optimal operating conditions for the linked bioreactors wouldbe known before the start of the experiment, it would likely be mostefficient to immediately direct the cell bleed (from the first day thecell bleed is started) to the production (CSTR) reactor.

Once the cell bleed from the N-1 perfusion reactors was directed intothe production bioreactors (CSTR's), the cell density of those reactorsvery quickly increased to above 20×10⁶ viable cells/mL (FIG. 14). Theflowrate of concentrated feeds added directly to the productionbioreactors was increased as necessary as the cell density increased inorder to maintain sufficient nutrients for continued cell growth andprotein production. The nutrients were monitored by amino acid analysis(HPLC) and other metabolites (lactate, glucose, ammonia, osmoticstrength, etc.) were monitored by means of NovaFlex Analyzers (NovaBiomedical, Waltham, Mass.). The feeding rates were adjusted so that thesum of the residual amino acids (not including alanine, as alanine isoften at high levels in culture as a metabolic byproduct) was maintainedabove 30 millimolar and that any particular amino acid was not less than0.2 millimolar. The feed consisted of a concentrated solution of glucose(50 grams/L), amino acids (600 millimolar), shear protectant (polyvinylalcohol at 5.12 grams/L), vitamins and trace elements (some detailed intable 6).

TABLE 6 Media and feeds composition details Total Amino Glu- Sodium-L-Osmotic Acids cose Lactate Strength Medium (mM) (g/L) (g/L) (mOsm) Basalmedium for N-1 perfusion 120 4 0 280 and CSTR production bioreactorsPerfusion medium (days 1-43) 60 4.2 2.1 319 Concentrated feed medium for607 50 0 ~1100 production (CSTR) bioreactors Saline diluent solutionused for 0 0 0 250 achieving high dilution rate in productionbioreactors from days 10 to 26 (20 mM KCl, balance NaCl) Concentratedglucose solution 0 500 0 ~2777 used for production bioreactors

The dipeptide glycine-tyrosine was also used in the concentrated feed toreduce complications of tyrosine solubility. For every 100 mL ofconcentrated feed medium added to the production bioreactor 2.5 mL of a400 millimolar acidic stock solution of cystine was also added.

Since the goal of the experiment was to explore different dilution ratesin the production bioreactor, it was realized that a saline diluentsolution is also necessary to be fed to the production bioreactors,particularly when high dilution rates were being explored. This salinediluent consisted of water with 20 millimolar potassium chloride and thebalance sodium chloride with a final osmotic strength of 250 mOsm/kg.Thus, in an embodiment, the production bioreactor is fed with a salinesolution to help adjust the dilution rate and also to control theosmotic strength of the CSTR production bioreactor. In anotherembodiment, the CSTR production bioreactor is fed with a saline solutionin an amount sufficient for the medium in the production bioreactor toachieve a final osmotic strength that is optimal for an increasedproductivity of the cells. In another embodiment, the CSTR productionbioreactor is fed with a saline solution in an amount sufficient for themedium in the production bioreactor to achieve a final osmotic strengthof about 100 to about 500 mOsm/kg, or to achieve a final osmoticstrength of about 150 to about 400 mOsm/kg, or to achieve a finalosmotic strength of about 150 to about 350 mOsm/kg, or any specificvalue therebetween. In another embodiment, the saline solution added tothe production bioreactor is sufficient to produce a final osmoticstrength of about 250 mOsm/kg.

Initially the levels of lactate in the production bioreactors increasedvery quickly (days 8-10, FIG. 11). For this reason, once the glucose wasexhausted in the production bioreactors (day 10, FIG. 10) the feeding ofglucose to the production bioreactors was limited. The glucose level wascontrolled by using a continuous feed of concentrated nutrients andglucose that was at a level of glucose that was predicted to be belowthat needed to balance the amount of amino acids contained in the feed(details in table 6 and as previously described). An additional feed ofconcentrated 500 gram/L glucose was fed directly and slowly to the twoproduction bioreactors by means of pumps that were activated when pHreached the high-end set-point of 7.125. This method of control oflactate, High-end pH-Controlled Delivery of Glucose (HIPDOG) isdescribed in detail in Gagnon et. al. 2011, Biotechnol Bioeng108:1328-37. The HIPDOG control was used until day 26 in the productionbioreactor using cell bleed from the target 40×10⁶ N-1 perfusionbioreactor, and until day 22 in the production bioreactor using cellbleed from the target 80×10⁶ N-1 perfusion bioreactor. After thesetime-points, day 22 and day 26, the residual glucose levels in theproduction bioreactors were maintained in what is believed to be anon-limiting range, generally above 0.3 grams/L, but more often in the1-3 gram/L range (FIG. 10) for the remainder of the experiments bycontinuous feeding of the concentrated glucose solution (500 gram/L).

FIG. 14 shows the viable cell densities in the production bioreactors.The cell densities quickly increased initially, but then began tostabilize by day 10 to 14. The volume and concentration of cellsentering the production bioreactors from the N-1 perfusion reactors wereheld constant from about day 10 until about day 32. At that point thevolume of cells entering the production bioreactors was cut in half inorder to simulate a larger bioreactor volume ratio difference of 1:10(N-1 compared to production bioreactor volume). The ratio was changedagain on about day 43 to simulate a ratio of 1:20. The timing of thesechanges are indicated with text on the graph of cell density in theproduction (CSTR) bioreactors (FIG. 7).

From day 13 to day 18 the volume of nutrient feeds to both of theproduction bioreactors were held roughly constant. The dilution rate(measured in reactor volumes per day, or reciprocal time) of theproduction bioreactor (shown in FIG. 15 and also in text on FIGS. 14 and17) is the sum of the volumes of cell bleed, concentrated nutrient feedand saline diluent, concentrated glucose feed, and any minor additionsof anti-foam and base titrant to control pH on the low end, divided bythe reactor volume. The dilution rate was held relatively constant at anaverage value of 0.64 reactor volumes per day (FIG. 15) from days 13-18for both production bioreactors. During this time period when thedilution rate on the two production bioreactors was held roughlyconstant, the cell density (FIG. 14), cell viability (FIG. 16), titer orproduct concentration (FIG. 17), and most other metabolites reachednearly constant values indicating that a steady-state condition had beenachieved.

After day 18 the dilution rates of the production bioreactors wereslowly reduced over the course of several days to 0.30 reactor volumesper day for the culture receiving cell bleed from the target 40×10⁶viable cell/mL N-1 perfusion reactor, and 0.33 reactor volumes per dayfor the culture receiving cell bleed from the target 80×10⁶ viablecell/mL N-1 perfusion reactor. The reduction in dilution rates wasaccomplished by slowly decreasing the volume ratio of saline diluent toconcentrated feed that was being added directly to the productionbioreactors. As before, the cultures were monitored for nutrient levelsand the concentrated nutrient feed rate was adjusted as necessary tofeed sufficient nutrients for the changing cell densities.

By day 21 both production bioreactors had reached the next targeteddilution rate (0.30 and 0.33 reactor volumes/day). The productionbioreactors were held at the target dilution rates for another 5 daysuntil most culture parameters had reached nearly constant values. Due tothe lower dilution rates that were being tested at this point, and forthe rest of the experiment, it was impractical to wait until thecultures had reached completely unchanging values for every parameter.

From day 26 to 28 the dilution rates of the production bioreactors wereagain decreased by decreasing the volume ratio of saline diluent toconcentrated feed that was being added directly to the productionbioreactors. To achieve the reduced dilution rates that were heldconstant from days 28-31 (0.21 and 0.25 reactor volumes/day as shown inthe figures, e.g., 15) it was necessary to completely eliminate the useof the saline diluent. From day 28 forward the feeds to the productionbioreactors consisted only of the cell bleed from the associated N-1perfusion reactor, the concentrated nutrient feed, a concentratedglucose feed, and trace amounts of antifoam. Negligible amounts of basetitrant were required for the remainder of the experiment since thecells had shifted into a metabolic state in which no net lactic acid wasbeing produced. Again the majority of metabolic parameters stabilized tonear constant values by day 31.

On day 31 the volume of cell bleeds from the N-1 perfusion bioreactorsbeing added to their associated production bioreactor (CSTR) werereduced by one half in order to simulate a production bioreactor with avolume ten times that of the N-1 reactor. To restate this, the volume ofcell bleed being removed from the N-1 perfusion bioreactor was notchanged, but a larger fraction of that cell bleed was now sent to wasterather than being added to the associated production bioreactor. Thischange in simulated volume ratio was necessary to attempt to study adilution rate of the production bioreactor lower than that alreadytested without limiting the availability of nutrients to the culture.After the change in simulated volumes, the viable cell density of theproduction bioreactor receiving cell bleed from the target 80×10⁶cell/mL N-1 perfusion bioreactor decreased to just over 50×10⁶ cells/mL.As this occurred, the requirement for concentrated nutrient feedslightly decreased until the effective dilution rate of the culturereached a value of approximately 0.15 reactor volumes/day on day 34.This dilution rate was held constant until day 41 at which time theexperiment was completed for the production bioreactor receiving cellbleed from the target 80×10⁶ cell/mL N-1 perfusion bioreactor. In asimilar manner the effective dilution rate of the production bioreactorreceiving cell bleed from the target 40×10⁶ cell/mL N-1 perfusionbioreactor decreased to about 0.14 reactor volumes/day on day 34 and washeld there until day 43 when the majority of parameters reachedsteady-state values.

On day 43 the simulated volume ratio of N-1 perfusion bioreactor toproduction bioreactor was changed again, this time to 1:20. This allowedfor the exploration of a dilution rate of 0.09 reactor volumes/day fromdays 46 to 50 for the production bioreactor receiving cell bleed fromthe target 40×10⁶ cell/mL N-1 perfusion bioreactor (FIG. 15). Also onday 43, the formulation of perfusion medium being added to the target40×10⁶ cell/mL N-1 perfusion bioreactor was slightly altered. TheL-sodium lactate level in the perfusion medium prior to day 43 had been2.1 grams/L. On day 43 a switch was made to use a perfusion mediumcontaining sodium DL-lactate, a more commercially available andeconomical alternative to the pure L-sodium lactate chemical. Anassumption was made that the sodium DL-lactate concentrated solutionobtained contained an equal ratio of D and L-isomers of lactate and thatthe cells in culture would likely only take up the L-isomer form. Thesodium DL-lactate was added to a level of 4.2 grams/L such that themolar concentration of L-lactate in the perfusion medium would still beapproximately 18.8 millimolar. An attempt was made to prepare theperfusion medium before and after the change to DL-lactate with the sameosmotic strength. This was accomplished by compensating for theadditional osmotic contribution of the added sodium D-lactate byreducing the amount of sodium chloride added to the perfusion medium. Asa final step prior to sterile filtration during the medium preparationprocess, the osmotic strength is measured by a freezing point osmometer.Sodium chloride is normally added to the medium preparation to adjustthe osmolality of the solution close to the desired value, in this caseapproximately 319 mOsm/kg. Despite the attempt to match the osmoticstrength of the perfusion media there was a slight increase in thesteady-state osmotic strength of the N-1 bioreactor around day 43 soonafter the switch to the perfusion medium containing the DL-lactate (FIG.18).

In separate flask tests of a pure form of sodium-D-lactate it wasdetermined that within the normal ranges of use (0-5 grams/L) theD-lactate form does not register with the analytical equipment using thestandard analysis methods. For this reason it is not surprising thatafter the change to the perfusion medium containing the DL-lactate (day43) the level of lactate as measured in the target 40×10⁶ cell/mL N-1perfusion bioreactor did not change appreciably (FIG. 11).

The volumetric productivity of the N-1 perfusion bioreactors operatingindependently as perfusion bioreactors are shown in FIG. 19. This figureassumes that the perfusion bioreactor is operating as a productionbioreactor and that the protein product is recovered from all streamsleaving the bioreactor (the cell bleed and the cell-free permeateleaving the cell retention system). The volumetric productivity alsoconsiders minor changes in product concentration in the bioreactor thatoccurred when a small amount of selective concentration of product inthe bioreactor occurred. This was mostly only evident in the N-1perfusion bioreactor with the target 80×10⁶ cell/mL density (FIG. 20,days 15-25).

Even though the operating conditions of the perfusion bioreactors arenot changed significantly after about day 10, and the cell density, cellviability, and other culture parameters are nearly unchanged from day 10to the end of the experiment for both perfusion bioreactors, thevolumetric productivity slowly declines by about 29% for the target40×10⁶ cell/mL N-1 perfusion bioreactor (from day 12 to 48) and about36% for the target 80×10⁶ cell/mL N-1 perfusion bioreactor (from day 12to 40). This loss in productivity is likely a result of moderate geneticinstability of the producer CHO cell line B.

The steady-state volumetric productivity of the combined N-1 perfusionlinked to production bioreactor (CSTR) system is shown in FIG. 21. Thevolumetric productivity is graphed relative to the dilution rates usedin the production bioreactors. The steady-state volumetricproductivities of the linked N-1 perfusion and CSTR productionbioreactor were calculated in the following manner. A material balancewas performed to determine the rate of production of product protein inthe production bioreactor. The material balance considered the productconcentration and volume of liquid entering the production bioreactorfrom the N-1 bioreactor, the rate of product concentration increase inthe production bioreactor, and the product concentration at anytime-point in the production bioreactor and its rate of removal (thedilution rate). This calculation yielded the total mass of productproduced in the CSTR which was then plotted against time. A linearregression was then used to more accurately determine the production ofmass per time per reactor volume value, only considering time pointswhen the dilution rate was held constant and the cell density hadreached a relatively constant value. This estimation was then added tothe production rate of the N-1 bioreactor (again assuming the correctsimulated volume ratio of N-1 to production bioreactor, and onlyconsidering the volume of fluid that enters the production bioreactor)for the final volumetric productivity estimation of the linked N-1perfusion bioreactor to production CSTR system for any particulardilution rate.

Multiplying the product concentration in the CSTR production bioreactorat the last time-point for any particular steady-state by the dilutionrate (in reciprocal time) will also yield an estimate of the overallvolumetric productivity of the linked bioreactor system. However, unlessthe product concentration is unchanging at the time-point used, thiscalculation will give a slight underestimate of the actual steady-statevolumetric productivity of the system.

As can be seen in FIG. 21, for many of the steady-state conditions, thevolumetric productivity of the combined N-1 perfusion and productionbioreactor is considerably higher than the volumetric productivities ofthe N-1 perfusion bioreactors operating independently. The one exceptionto this generalization might be the volumetric productivity of thetarget 80×10⁶ cell/mL N-1 perfusion bioreactor at the earliesttime-points on FIG. 19. This perfusion bioreactor was operating at aperfusion rate of 3.6 reactor volumes/day and a very high cell density.The practical operation of such a bioreactor at very large scale mightentail significant challenges due to the volumes of medium that would berequired to be prepared and stored, and also the difficulty of designinga cell retention system that could function at very large volumes.Additionally, since delivering sufficient oxygen and removing sufficientcarbon dioxide becomes more problematic at large scale, the very highviable cell densities (80×10⁶ cells/mL) might pose substantialengineering complications.

The linked N-1 perfusion to CSTR production bioreactor alleviates manyof these issues by dramatically reducing the size of the perfusionbioreactor and drastically reducing the overall volumes of mediumrequired. The linked bioreactor system maintains volumetricproductivities approximately the same as a continuous perfusionbioreactor operating at high density and at the high perfusion ratesnecessary to maintain that high cell density. FIG. 21 shows that theoptimal dilution rate for operating the CSTR production bioreactoroccurs over a quite broad range from around 0.1 to approximately 0.35reactor volumes per day. Due to the slow loss in productivity of thecell line as indicated in FIG. 19, and the fact that the high dilutionrate conditions were investigated first during the experiment, it islikely that the steady-state volumetric productivities as presented inFIG. 21 for the low dilution rates might be slight underestimates of theactual productivity achievable.

Table 7 lists some additional details that relate to the highestvolumetric productivity conditions for the two linked N-1 perfusion toCSTR systems as indicated on FIG. 21.

TABLE 7 Various process parameters relating to the two higheststeady-state volumetric productivities for the linked N-1 perfusion andCSTR production bioreactor system. A comparison of the perfusionbioreactors operating independently as production bioreactors and aspart of the linked bioreactor system. Maximum Total Medium Grams ofprotein Volumetric Dilution Rate Consumption Rate of product producedper Productivity of CSTR Linked Bioreactor liter of medium and (grams/LPerfusion rate (reactor System nutrient feed production (reactorvolumes/day (production bioreactor consumed bioreactor Processvolumes/day) or 1/day) volumes/day) (grams/L) volume/day) SustainablePerfusion with target 1.76 NA NA 0.24 0.42 40 × 10⁶ cells/mL (perfusionreactor operating independently) Linked N-1 Perfusion to CSTR NA 0.140.24 3.83 0.92 N-1 1/10th volume of CSTR (using bleed from target 40 ×10⁶ cells/mL perfusion reactor) Sustainable Perfusion with target 3.60NA NA 0.27 0.96 80 × 10⁶ cells/mL (perfusion reactor operatingindependently) Linked N-1 Perfusion to CSTR NA 0.25 0.82 1.34 1.10 N-1⅕th volume of CSTR (using bleed from target 80 × 10⁶ cells/mL perfusionreactor)

The efficiency of medium or feed use is significantly higher in thelinked bioreactor systems when compared with the perfusion bioreactorsoperating independently. The condition with the highest efficiency ofmedium usage, 3.83 grams product protein produced per liter of medium orfeed used in the system as a whole for the linked system, is quite high,even when compared with a mammalian cell culture bioreactor operating ina fed-batch mode.

While in the current experiment the cell bleed rate was adjustedmanually each day to ensure that the two N-1 perfusion culturesmaintained a cell density close to their target values, in an industrialsetting with a well understood process, the cell density is monitoredwith any number of sterilizable probes that utilize capacitancemeasurements (or other technologies) to estimate the viable cell mass inan operating bioreactor. A computer control algorithm could then bewritten to continuously vary the amount of cell bleed removed from theN-1 perfusion bioreactor. Presumably in a linked N-1 perfusion to CSTRproduction bioreactor system all cell bleed would be immediatelytransferred to the production bioreactor.

Since the effective dilution rate of the CSTR production bioreactor isdependent to a large degree on the volume of liquid entering the vesselfrom the continuous N-1 perfusion bioreactor, in an embodiment, the cellbleed from the N-1 perfusion bioreactor is further concentrated (byremoving additional liquid from the cell bleed) before it is transferredinto the CSTR production bioreactor. This concentration of the cellbleed is particularly valuable when the N-1 perfusion bioreactor isoperating at a low cell density and/or when it is determined inexperiments that low dilution rates in the CSTR production bioreactoryielded higher volumetric productivities for the system as a whole. Thefurther concentration of the cell bleed (if necessary) is performed byany practical cell retention method (continuous micro-filtration,acoustic wave settling, hydrocylone, etc.).

Since the cell bleed from the N-1 perfusion reactors brings aconsiderable fraction of the total volume of fluid into the productionreactor (depending upon the conditions), in an embodiment, to minimizethe addition of feed medium to the production bioreactor, a very highnutrient content medium is used to perfuse the N-1 bioreactor. Since thecells in the N-1 perfusion reactor would not be able to consume all ofthe nutrients from the perfusion medium, significant amounts ofnutrients would then flow in the production bioreactor. There is a limitto this capability however in that the growth rate of cells in thepresence of very high levels of some amino acids may be considerablylimited. Additionally, since most of the perfusion medium entering theN-1 bioreactor is lost to waste in the cell free permeate, theadditional costs of discarding large volumes of high nutrient contentperfusion medium would have to be considered.

In a similar manner to the reasoning for potentially concentrating thecell bleed before transfer, having a nutrient feed directly to the CSTRproduction bioreactor that is highly concentrated is important for thelinked system to function effectively and achieve high volumetricproductivities. A very dilute feed medium might dilute the biomass inthe CSTR production bioreactor by increasing the effective dilution rateof the bioreactor. With mammalian cells that typically have maximumdoubling times near or above one day, any effective dilution rate nearor above 1.0 reactor volumes/day (or 1/day) might result in a very lowcell density in the CSTR production bioreactor and therefore lowproductivity in the linked bioreactor system.

In the example we have described above, during conditions in whichdilution rates were above 0.20 reactor volumes/day (1/day) a salinediluent as described in table 3 was used to dilute the concentratedfeeds to the CSTR production bioreactor. FIG. 21 shows that reasonablevolumetric productivities can be achieved with the linked bioreactorsystem at dilution rates even as high as 0.64 reactor volumes per day.High dilution rates dilute the viable cell mass in the CSTR and tend todrive productivity lower, but high dilution rates also simultaneouslyflush cell-generated inhibitory compounds out of the bioreactor systemand therefore increase the growth rate of cells. These competing factorspartially balance each other out which allows for the cell density athigh dilution rates to still be maintained reasonably high, and as aresult reasonably high volumetric productivities can still be achieved.Thus, in one embodiment, the dilution rate of the production bioreactoris from about 0.1 volumes/day to about 1.25 volumes/day. In anotherembodiment, the dilution rate of the production bioreactor of thesubject technology is from about 0.2 volumes/day to about 1.0volumes/day; or from about 0.2 to about 0.75 volumes/day; or about 0.2volumes/day or about 0.65 volumes/day. In another embodiment, thedilution rate of the production bioreactor of the subject technology isabout 0.1 volumes/day; or is about 0.2 volumes/day; or is about 0.25volumes/day; or is about 0.3 volumes/day; or is about 0.35 volumes/day;or is about 0.4 volumes/day; or is about 0.45 volumes/day; or is about0.5 volumes/day; or is about 0.6 volumes/day; or is about 0.7volumes/day; or is about 0.8 volumes/day; or is about 1.0 volumes/day;or is about 1.25 volumes/day.

At 0.64 reactor volumes per day the average residence time of productbeing generated in the bioreactor is approximately 1.56 days. Thisresidence time is far less than the average residence time of productmade in a fed-batch bioreactor. For this reason, the linked N-1perfusion to CSTR production bioreactor system might have significantadvantages over the fed-batch mode of bioreactor operations in the caseof labile proteins. Additionally, through use of increased volumes ofthe saline diluent added to the production bioreactor the dilution rateof the linked bioreactor system could be manipulated to reduce theresidence time of product protein without overly reducing the volumetricproductivity of the system as a whole. The same result of course couldbe achieved by simply manipulating the concentration of the feed mediathat flows directly in to the production bioreactor. Thus, in anembodiment, the product residence time in the production bioreactor ofthe subject technology is from about 1 day to about 10 days. In anotherembodiment, the product residence time in the production bioreactor ofthe subject technology is about 1 day, or about 2.0 days; or about 2.5days; or about 3.0 days; or about 3.5 days; or about 4 days, or about 5days, or about 6 days, or about 7 days, or about 8 days, or about 9days, or about 10 days. In another embodiment, the residence time of thecells in the production bioreactor of the subject technology is lessthan about 10 days.

As mentioned earlier in this text, the strategy of limiting glucoseavailability in the CSTR production bioreactors was used at thebeginning of the experiment (HIPDOG technology). This limited theaccumulation of lactate in the bioreactors. Since the low dilution ratesfor the CSTR's proved to be the most productive, in an industrialapplication it might be of value to immediately or very quickly move toa condition of low dilution rates. Presumably when the first cells flowinto the production bioreactor CSTR from the N-1 perfusion bioreactorthe CSTR bioreactor will be a least partially full of fresh basalmedium. In a highly proliferative state in the presence of freelyavailable glucose those cells might start to produce large amounts oflactic acid, and with pH control this would also mean a high osmoticstrength as the lactic acid was neutralized. Any lactate accumulated(and its associated high osmotic strength) in the CSTR early in theculture while the bioreactor is filling or produced immediatelythereafter while at very low dilution rates would be very slow to bediluted out of the culture system. For this reason there is significantvalue in utilizing a glucose limiting technology such as high-endpH-controlled delivery of glucose (HIPDOG) in the initial startup phaseof the CSTR production bioreactor. Without some method of control ofinitial lactate production, high levels of lactate could accumulate inthe production bioreactor such that after only several days of operationlittle or no growth would be achievable in the production bioreactor.When CHO cells in culture are exposed to a high lactate concentrationsometimes a positive feed-back condition occurs in which lactateproduction from glucose is significantly upregulated. Such a conditionmight significantly delay or even make unachievable the attainment of asteady-state condition with high volumetric productivity at low dilutionrates. While the HIPDOG glucose limiting strategy worked well in thecurrent experiment, a shift to a low pH control set-point or temperaturelower than 36.5° C. might also or alternatively be used to minimizelactate formation in the CSTR production bioreactor for a period of timewhen it is first started. Both lower pH and lower temperature might slowgrowth and decrease cellular productivity, so the use of these setpointcontrol changes would have to be balanced against the need to controllactate formation. The optimal pH setpoint would likely be lower than7.0 for this purpose, likely between 6.7 and 7.0. The optimaltemperature would likely be between 30-35° C.

It will be apparent to those skilled in the art that variousmodifications and variations can be made to the described embodimentswithout departing from the spirit and scope of the claimed methods.Thus, it is intended that present claimed methods cover themodifications and variations of the embodiments described hereinprovided that they come within the scope of the appended claims andtheir equivalents.

INDUSTRIAL APPLICABILITY

The device and methods disclosed herein are useful forlinked-perfusion-to-CSTR cell culture biomaufacturing, and thus forimproving industrial methods for manufacturing recombinant and/ortherapeutic proteins.

1. A method of producing a protein of interest, comprising: (a)culturing cells comprising a gene that encodes the protein of interestin a culture bioreactor (N-1 bioreactor); (b) inoculating a productionbioreactor (N bioreactor) with cells obtained from step (a); and (c)culturing the cells in the production bioreactor under conditions thatallow production of the protein of interest.
 2. The method according toclaim 1, wherein the method further comprises step (d) harvesting theprotein of interest from the production bioreactor.
 3. The methodaccording to claim 1, wherein the culture bioreactor is a continuousperfusion culture bioreactor and the production bioreactor is acontinuously stirred tank reactor (CSTR) production bioreactor.
 4. Themethod according to claim 1, wherein the production bioreactor has nocell retention device.
 5. The method according to claim 1, whereinvolume ratio of the culture bioreactor to the production bioreactor isabout 1:1 to about 1:20 or about 1:1 to about 1:5 or about 1:5.
 6. Themethod according to claim 1, wherein the inoculation in step (b) is bytransferring cells from the culture bioreactor to the productionbioreactor.
 7. The method according to claim 6, wherein the celltransfer is by cell bleed in continuous or semi-continuous modes.
 8. Themethod according to claim 7, wherein the cell transfer is insemi-continuous mode comprising the cell transfer once at every periodof time from 2 minutes to 24 hours or at any interval therebetween. 9.The method according to claim 1, wherein step (a) optionally alternatesbetween a first and second culture bioreactors to allow for renewal andcontinuous production of culture cells for use in step (b).
 10. Themethod according to claim 9, wherein the second culture bioreactor is acontinuous perfusion culture bioreactor.
 11. The method according toclaim 1, wherein the production bioreactor operates continuously for aperiod of greater than 3 weeks or for a period of greater than 4 weeksor for a period of greater than 5 weeks or for a period of greater than6 weeks.
 12. The method according to claim 1, wherein harvesting step(d) is continuous.
 13. The method according to claim 1, wherein thecells are CHO cells, HEK-293 cells, VERO cells, NSO cells, PER.C6 cells,Sp2/0 cells, BHK cells, MDCK cells, MDBK cells or COS cells.
 14. Themethod according to claim 1, wherein the production bioreactor has avolumetric productivity of at least 0.6 grams per liter per day for aperiod of at least 14 days or for a period of at least 20 days or for aperiod of at least 30 days.
 15. The method according to claim 1, whereinthe production bioreactor has a product residence time of about 1 toabout 10 days.
 16. The method according to claim 1, wherein theproduction bioreactor has a dilution rate of about 1 to about 0.1 volumeper day.
 17. A linked culture process, comprising: (a) culturing cellscomprising a gene that encodes the protein of interest in a culturebioreactor (N-1 bioreactor); (b) inoculating a production bioreactor (Nbioreactor) with cells obtained from step (a); and (c) culturing thecells in the production bioreactor under conditions that allowproduction of the protein of interest.
 18. The linked culture processaccording to claim 17, wherein the linked culture process furthercomprises step (d) harvesting the protein of interest from theproduction bioreactor.
 19. The linked culture process according to claim17, wherein the culture bioreactor is a continuous perfusion culturebioreactor and the continuous production bioreactor is a continuouslystirred tank reactor (CSTR) production bioreactor.
 20. The linkedculture process according to claim 17, wherein the production bioreactorhas no cell retention device. 21-61. (canceled)